High octane unleaded gasoline production

ABSTRACT

A HIGH OCTANE, UNLEADED GASOLINE POOL IS PRODUCED VIA AN INTEGRATED REFINERY OPERTION. THE INDIVIDUAL PROCESSES INTEGRATED IN COMBINATION WITH EACH OTHER RE HYDROCRACKING, LOW SEVERITY CATALYTIC REFORMING, SATURATE CRACKING AND ALKYLATION. CATALYTIC REFORMING IS EFFECTED AT LOW SEVERITY CONDITIONS TO MAXIMIZE THE REACTION OF NAPHTHENE DEHYDROGENATION TO PRODUCE AROMATICS WHICH SIMULTANEOUSLY INHIBITING HE DEHYDROCYCLIZATION AND CRACKING OF PARAFFINIC HYDROCARBONS.

March 21, 1972 P. SCHULLER HIGH OCTANE UNLEADED GASOLINE PRODUCTIONFiled July 10, 1970 m qumum m um IN VEN TOR Richard P. Schul/er mc us mS m mum mqs mmc mum 9 5 mm cotu lmEm 3 A TTOR/VEYS 3,650,943 HIGH OCTANEUNLEADED GASOLINE PRODUCTION Richard P. Schuller, Des Plaines, Ill.,assignor to Universal Oil Products Company, Des Plaines, Ill. Filed July10, 1970, Ser. No. 53,719 Int. Cl. C10g 13/00, 21/00; C07g 3/54 US. Cl.208-60 10 Claims ABSTRACT OF THE DISCLOSURE APPLICABILITY OF INVENTIONThe present invention relates to the conversion of heavyhydrocarbonaceous charge stocks into lower-boiling hydrocarbon products.More specifically, the inventive concept herein described is directedtoward an integrated refinery process for producing a high octane,unleaded gasoline pool. Relatively recent investigations into the causesand cures of environmental pollution have indicated that more than about50.0% of Violence to the atmosphere stems from vehicular exhaustconsisting primarily of unburned hydrocarbons and carbon monoxide. Theseinvestigations have brought about the development of catalyticconverters which, when installed in the automotive exhaust system, arecapable of converting more than 90.0% of the noxious components intoinnocuous material prior to the discharge thereof into the atmosphere.In developing these catalytic converters, it was learned that theefiiciency of conversion and the stability of the catalytic compositewere severely impaired when the exhaust fumes resulted from thecombustion of leadcontaining motor fuels. Compared to operations of thecatalytic converter during the combustion of clear, unleaded gasolines,both the conversion of noxious components and catalyst stabilitydecreased as much as 50.0%. Therefore, it has been recognized throughoutthe petroleum industry, as well as in the major gasoline-consumingcountries, that suitable gasoline must be produced for consumption ininternal combustion engines without requiring the addition of lead toincrease the octane rating and enhance the anti-knock propertiesthereof. Also, being recognized is the fact that the unburnedhydrocarbons and carbon monoxide are not the only extremely dangerouspollutants being discharged via vehicular exhaust. Japan has experiencedan increase in the incidence of lead poisoning, and has enactedlegislation to reduce significantly the quantity of lead in motor fuelgasolines.

One natural consequence of the removal of lead from motor fuel gasoline,in addition to many others including the possibility of future enginemodifications, resides in the fact that petroleum refining operationsmust be changed in order to produce voluminous quantities of highoctane, unleaded motor fuels. 'One well known refining process capableof significantly improving the octane rating of gasoline boiling rangefractions is the catalytic reforming process. In such a process theprimary octaneimproving reactions are naphthene dehydrogenation,naphthene dehydroisomerization and parafiin dehydrocyclization.Naphthene dehydrogenation is extremely rapid, and constitutes theprincipal octane-improving renited States Patent action. With respect tofive-membered alkyl naphthenes, it is necessary to effect isomerizationto produce a six-membered ring naphthene followed by dehydrogenation toan aromatic hydrocarbon. Parafiin aromatization is achieved throughdehydrocyclization of straight-chain paraffins having at least sixcarbon atoms per molecule. The latter reaction is limited in catalyticreforming operations since the aromatic concentration increases throughthe reforming zone, thereby decreasing the rate of additionaldehydrocyclization. Unreacted, relatively low octane paraflins are,therefore, present in the catalytically reformed effluent, andeifectively reduce the overall octane rating thereof. At a relativelyhigh severity, the paraffinic hydrocarbons in the reforming zone aresubjected to cracking. While this partially increases the octane ratingof gasoline boiling range material, substantial quantities of normallygaseous material are produced. In view of the fact that hydrogen ispresent within the reaction zone, the light gaseous material issubstantially completely saturated and comprises principally methane,ethane, propane and butane.

At a relatively low reforming severity, paraflin cracking is decreasedwith the result that an increased quantity of low octane ratingsaturates are produced. In order to upgrade the overall quality of thegasoline pool, either the addition of lead becomes necessary, or the lowoctane rating saturates must be subjected to further processing. Ashereinabove set forth, subsequent processing of the saturates for octanerating improvement can be eliminated by increasing the operatingseverity within the catalytic reforming reaction zone. A high severityoperation produces a two-fold effect while increasing the octane rating;first, additional high octane aromatic components are produced and,secondly, the low octane rating components are at least partiallyeliminated through conversion into either aromatic components, or lightnormally gaseous hydrocarbons. The results, therefore, include lowerliquid yields of gasoline due both to shrinkage in molecular size whenparaflins and naphthenes are converted to aromatics, and to theproduction of the aforesaid light gaseous components. These problems arefurther compounded when the desired end result is the production of ahigh octane, unleaded gasoline pool. In accordance with the integratedrefinery operations constituting the present invention, 'a low severitycatalytic reforming unit is dove-tailed with at least a hydrocrackingunit, a saturate cracking unit, and an alkylation unit. As hereinafterindicated, the end result is the production of a high octane, unleadedgasoline pool unaccompanied by substantial liquid yield loss.

The hydrocarbonaceous charge stocks, contemplated for conversion inaccordance with the present invention, constitute heavier-than-gasolinehydrocarbon fractions and or distillates. Since gasoline boiling rangehydrocarbons is intended to connote those hydrocarbons having an initialboiling point of about F. to about 125 F. and an end boiling point fromabout 400 -F. to about 450 F., the contemplated charge stocks will haveinitial boiling points above about 400 F. The end boiling point of thecharge stocks will be about 1050 F., or less, generally considered to bethat temperature at which distillation can be eifected without incurringthermal cracking. With respect to those hydrocarbonaceous materialscontaining hydrocarbons which would normally boil above a temperature 0F. (considered in the art as black oils) they are not considered for useuntil the 1050 F.-plus material, constituting primarily non-distillableasphaltenes, has been removed. Thus, suitable charge stocks intended forconversion through the use of the present process include kerosenefractions, light gas oils boiling up to the temperatures of about 600F., heavy vacuum or atmospheric gas oils boiling up to a temperature ofabout 1050 P.

and either intermediate, or overlapping fractions and mixtures thereof.With respect to the gasoline boiling'range hydrocarbons, light naphthagenerally refers to a hydrocarbon mixture concentrated in hydrocarbonshaving five and six carbons atoms per molecule. Light naphthas can berecovered directly from a crude distillation unit and have an endboiling point in the range of about 175 F. to about 200 F. The heavynaphtha is considered a hydrocarbon mixture having an initial boilingpoint of about 180 F. and end boiling point of about 400 F. to about 450F., and includes primarily those hydrocarbons having seven or morecarbon atoms per molecule.

OBJECTS AND EMBODIMENTS A principal object of the present invention isto convert heavier-than-gasoline hydrocarbon distillates intolowerboiling, gasoline boiling range hydrocarbon products. A corollaryobjective resides in the production of a high octane, unleaded motorfuel gasoline pool.

Another object of my invention is to provide an integrated refineryoperation for producing high liquid yields of a high octane, unleadedgasoline pool.

In one embodiment, therefore, my invention affords a process forproducing a high octane, unleaded gasoline pool which comprises thesteps of: (a) reacting a heavierthan-gasoline charge stock with hydrogenin a catalytic hydrocracking reaction zone, at hydrocracking conditionsselected to produce gasoline boiling range hydrocarbon products; (b)separating the resulting hydrocracked product efiluent to provide afirst substantially saturated vaporous phase and a gasoline boilingrange, normally liquid stream; (c) reacting at least a portion of saidliquid stream and hydrogen in a low-severity catalytic reformingreaction zone, at reforming conditions selected to convert naphthenichydrocarbons into aromatic hydrocarbons; (d) separating the resultingreformed product effiuent to provide an aromatic concentrate, asaturated normally liquid stream and a second substantially saturatedvapor phase; (e) reacting at least a portion of said saturated normallyliquid stream in a saturate cracking reaction zone, at crackingconditions selected to produce a cracked gasoline boiling range liquidstream and a substantially unsaturated vaporous phase; (f) reacting atleast a portion of said unsaturated vaporous phase with at least aportion of said first and second saturated vaporous phases in analkylation reaction zone, at alkylating conditions selected to producean alkylate gasoline boiling range, normally liquid stream; and (g)recovering said aromatic concentrate, said cracked gasoline stream andsaid alkylate gasoline stream as said high octane, unleaded gasolinepool.

Other embodiments of my invention involve the use of various catalyticcomposites, operating conditions and processing techniques. In one ofsuch other embodiments, the first and second saturated vaporous phasesare sepa rated to provide a butane concentrate, at least a portion ofwhich is reacted with at least a portion of the unsaturated vaporousphase in said alkylation reaction zone. In another such embodiment, theunsaturated vaporous phase is separated to provide a propyleneconcentrate and a butylene concentrate, the latter being reacted, atleast in part, with at least a portion of said first and second vaporousphases in said alkylation reaction zone. In a particularly preferredembodiment, the butane concentrate is separated into a normal butaneconcentrate and an isobutane concentrate, the former being reacted withhydrogen in an isomerization reaction zone to produce isobutane isomers.These, as well as other objects and embodiments, will become evidentfrom the following more detailed description of the process encompassedby the present inventive concept.

SUMMARY OF THE INVENTION As hereinbefore set forth, the integratedrefinery process of the present invention incorporates a hydrocrackingzone, a catalytic reforming zone, a saturate cracking zone and analkylation reaction zone. Additionally, in other embodiments, theoverall process includes an isomerization reaction zone and a solventextraction zone. In order that a clear understanding of the integratedrefinery process of the present invention is obtained, a briefdescription of the various reaction zones and separation zones, utilizedin one or more embodiments of the process is believed to be warranted.In describing each individual zone, one or more references to UnitedStates Patents will be made in order that more detail will be readilyavailable where desired. Such references are not intended to beexhaustive or limiting, but merely exemplary.

HYDROCRACKING ZONE The principal function of the hydrocracking reactionzone is to convert heavier-than-gasoline components into lower boiling,normally liquid products boiling within the desired gasoline boilingrange. Depending upon the physical and chemical characteristics of thecharge to the hydrocracking zone, the desired reactions will be effectedin a single stage or in a multiple-stage system. It is generallyconceded, in the hydrocracking art, that the heavier fractions derivedfrom crude oils are contaminated by substantial quantities of bothsulfurous and nitrogenous compounds. Therefore, one stage of the overallhydrocracking system will be a hydrorefining reaction zone wherein thesulfurous and nitrogenous compounds are converted into hydrogen sulfide,ammonia and hydrocarbons, the latter being subsequently hydrocracked toform the lower-boiling gasoline components. Exemplary of thehydrocracking process are those schemes and techniques found in US.Pats. 3,252,018 (Cl. 208-59), 3,502,572 (Cl. 208-111), and 3,472,758(Cl. 208-59).

The hydrocracking reaction is generally effected at elevated pressureswithin the range of about 500 to about 5,000 p.s.i.g., and preferablyfrom 1,500 to about 3,000 p.s.i.g. Circulating hydrogen is admixed withthe charge to the hydrocracking reaction zone in an amount of about3,000 to about 50,000 scf./bbl., and more often in the range of about5,000 to about 20,000 scf./bbl. The hydrogen and charge stock contactsthe catalytic composite, disposed within the hydrocracking reactionzone, at a liquid hourly space velocity of 0.25 to about 5.0, andpreferably from about 0.5 to about 3.0. Liquid hourly space velocity isdefined as the volumes of fresh feed charge stock per hour, computed at60 F., per volume of catalyst disposed within the reaction zone. Sincethe bulk of the reactions being effected are exothermic in nature, anincreasing temperature gradient will be experienced as the charge stocktraverses the catalyst bed. The maximum catalyst bed temperature isgenerally maintained in the range of 700 F. to about 900 F., and may becontrolled through the use of conventional quench streams which areintroduced at intermediate loci of the catalyst bed. Various componentsof the hydrocracked product effluent which boil above the desired endpoint of the gasoline product may be recycled to combine With the freshfeed charge stock for further conversion. With respect to the combinedliquid feed ratio, defined as total volumes of liquid feed per volume offresh feed, a range of 1.1 to about 6.0 is not uncommon, although it isgenerally preferred that the combined feed ratio be in the range ofabout 1.5 to about 3.0. The selection of any precise combination ofoperating variables depends on the characteristics of the charge stock,the desired product and the length of time the catalytic composite hasbeen in service.

The hydrocracking catalyst is a composite of a porous carrier materialand one or more catalytically active metallic components generallyselected from the metals of Groups V-B, VI-B and VIII of the PeriodicTable. With respect to the porous carrier material, it may be amorphous,or zeolitic in nature, the latter including well known crystallinealuminosilicates such as faujasite, mordenite, etc. Preferred carriermaterials are those comprising alumina and .silica, with the latterbeing in a concentration of about 10.0% to about 90.0% by weight.

The catalytically active metallic components are selected from the groupof metals consisting of vanadium, niobium, tantalum, chromium,molybdenum, tungsten, iron, cobalt, nickel, ruthenium, rhodium,palladium, osmium, iridium and platinum. The Group VI-B metals areutilized in a concentration of about 4.0% to about 30.0% by weight,while the Group V-B and iron-group metals are utilized in a lowerconcentration in the range of about 1.0% to about 10.0% by weight. Whenthe catalytic composite comprises a Group VHI noble metal, particularlyplatinum and or palladium, the concentration will be in the range ofabout 0.2 to about 2.0% by weight.

A multiple-stage hydrocracking system is preferred for utilization inthe present process. In the first stage, containing a catalyticcomposite of about 1.8% by weight of nickel and 16.0% by Weight ofmolybdenum, combined with a silica-alumina carrier material containing37.0% by weight of silica, sulfurous and nitrogenous compounds areremoved and some hydrocracking to lower-boiling products is effected. Inthe second stage, the preferred catalytic composite comprises about 5.0%by weight of nickel combined with a faujasitic carrier material, 90.0%by weight of which is zeolitic.

CATALYTIC REFORMING ZONE The charge to the catalytic reforming zone isgenerally derived from at least two sources. The greater proportion ofthe charge is the gasoline boiling range efiluent from the hydrocrackingreaction zone. A second source constitutes those naphtha fractionsderived from the original crude oil. Since the latter are generallycontaminated by sulfurous and nitrogenous compounds, the catalyticreforming reaction zone may have integrated therein a hydrorefining zonewhich necessarily treats only that portion of the charge derived fromthe original petroleum crude oil. Catalytic composites, for utilizationin the reforming reaction zone, include a refractory inorganic oxidecarrier containing a reactive metallic component generally selected fromthe noble metals of Group VIII. Recent developments in the area ofcatalytic reforming have indicated that catalyst activity and stabilityis enhanced through the addition of a Group VII-B or IV-A metalcomponent, particularly rhenium and/or germanium. Suitable porouscarrier materials include alumina, crystalline aluminosilicates such asthe faujasites, or mordenite, or combinations of alumina with thecrystalline alumino'silicates. Generally favored metallic componentsinclude ruthenium, rhodium, palladium, osmium, iridium, platinum,rhenium and germanium. These metalhc components are employed inconcentrations ranging from about 0.01% to about 2.0% by weight.Reforming catalysts may also contain combined halogen selected from thegroup of fluorine, chlorine, bromine, iodine and mixtures thereof.

Illustrations of catalytic reforming process schemes are found in US.Pats. 2,905,620 (Cl. 208-65), 3,000,812 (Cl. 208-138) and 3,296,118 (Cl.208-100). Effective reforming operating conditions include temperatureswithm the range of about 800 F. to about 1100 F., and preferably fromabout 850 to about 1050 F. The liquid hourly space velocity ispreferably in the range of about 1.0 to about 5.0, although spacevelocities from about 0.5 to about 15.0 may be employed. The quantity ofhydrogen-rich recycle gas in admixture with the hydrocarbon feed stockto the reforming reaction zones, is generally from about 1.0 to about20.0 moles of hydrogen per mole of hydrocarbon. The reforming reactionzone effluent is introduced into a high pressure separation zone at atemperature of about 60 F. to about 140 F., in order to separate lightercomponents from heavier, normally liquid components. Since normalreforming operations produce large quantities of hydrogen, a certainamount of the recycle gaseous stream is generally removed from thereforming system by way of pressure control. It is within the scope ofthe present invention that such excess hydrogen be employed in thehydrogen-consuming hydrocracking reaction zone as make-up hydrogen.Pressures in the range of about to about 1500 p.s.i.g. are suitable foreffecting catalytic reforming reactions.

With respect to the catalytic reforming reaction zone utilized in thepresent combination process, the reactions effected therein areconducted at a relatively low operating severity. To those familiar withthe catalytic reforming art, the term relatively high severity indicateshigh temperature or low space velocity, or both high temperature and lowspace velocity. The most noticeable direct result of high severityoperation is found in the octane rating of the normally liquid productefil'uent. While the reforming zone utilized in the present process doesnot necessarily upgrade the octane rating of the charge stock to thelevel ultimately attained with respect to the gasoline pool, the chargestock is substantially improved in octane rating.

In the present specification and accompanying claims, the term lowseverity reforming conditions alludes to a reforming process in whichsubstantial quantities of naphthenes are dehydrogenated to high octanearomatic compounds, while the dehydrocyclization and cracking ofparaflinic hydrocarbons are substantially inhibited. Low severityreforming operations may be defined by stating that from about 80.0 toabout 100.0 moles of aromatics are produced for every 100.0 moles ofnaphthenes in the charge stock, while less than about 40.0 moles ofaromatics are produced for every 100.0 moles of alkanes. In determiningthe degree of conversion of naphthenes to aromatics (dehydrogenation)and alkanes to aromatics (dehydrocyclization), it is assumed thatalrelatively small amount of naphthenes are cracked or otherwiseconverted to hydrocarbons other than aromatics, and that a major portionof the alkanes which disappear are converted to aromatic hydrocarbonswith some naphthenes and higher molecular weight alkanes being convertedto low molecular weight normally gaseous components.

SATURATE CRACKING ZONE As hereinafter set forth in greater detail, theproduct effluent from the catalytic reforming zone is separated toprovide an aromatic concentrate and a saturated liquid streamprincipally comprising paraffins and naphthenes. The primary function ofthe saturate cracking zone is to crack the saturated liquid stream. Thesaturate cracking zone may comprise a catalytic cracking system or athermal cracking unit. Although the preferred scheme utilizes means forseparately recovering an aromatic concentrate from the reforming zoneeffluent, it is understood that the entire effluent may be introducedinto the saturate cracking zone. In either situation, the sauratecracking zone must be capable of selectively cracking the feed stocksaturates to lower molecular weight hydrocarbons with the production ofdry gases such as methane, ethane, ethylene or acetylene beingminimized, whereas the production of propane, propylene, butanes,butylenes and cracked gasoline is maximized. Saturate cracking producescracked gasoline and valuable light hydrocarbons from the aromaticprecursors not converted in the reforming zone, due to the requirementof the low severity conditions, in order to obtain an overall advantagein liquid yield with respect to the high octane gasoline pool.

The materials produced in the saturate cracking zone, in addition to arelatively high octane cracking gasoline, include propane, propylene,normal and isobutane, normal and isobutene, pentanes, and pentenes.These products constitute excellent charge stocks for other processescapable of producing gasoline components such as amines, esters, ethers,ketones, branched-chain paraffins and alcohols. Preferably, the pentanesand heavier hydrocarbons are considered part of cracked gasoline and areintroduced as such into the high octane gasoline pool. The remainder ofthe olefinic portion of the hydrocarbons are especially suited forconversion to the previously described gasoline components, whereas theparaffinic portion of the saturate cracking zone effluent, containing arelatively large quantity of branched-chain isomers, is suited for theproduction of alkylate gasoline.

Typical, individual gasoline components which can be produced from thelight hydrocarbons emanating from the saturate cracking zone includemethyl alcohol, ethyl alcohol, isopropyl alcohol, isobutyl alcohol,tertiary butyl alcohol, iso-amyl alcohol, tertiary amyl alcohol,hexanol, isopropylamine, n-butylamine, diethylamine, tri-ethyl amine,methyl acetate, ethyl acetate, isopropyl acetate, isobutyl acetate,propylene oxide, n-propyl ether, isopropyl ether, iso-amyl ether,methyl-ethyl ketone, diethyl ketone, C -alkylate and C -alkylate. Inaccordance with the present invention, the butylenes and butanes areutilized as the feed stock to an alkylation reaction zone; the treatmentof the propylene concentrate involves either hydrolysis to produceisopropyl alcohol, or alkylation to produce a C alkylate.

The saturate cracking zone requires the use of high activity catalystsand elevated temperature. Preferred reaction temperatures lie within therange of about 850 F. to about 1200 F. An important operating parameterfor the selected production of large quantities of propylene andbutylene is the contact time between the saturated cracking zone feedand the catalyst contained therein. In fixed-bed cracking, whichgenerally incorporates a oncethrough operation, the weight ratio ofolefins to saturates is almost directly related to the space velocitybeing utilized with the reaction zone. Increasing the space velocity ofthe saturated feed increases the amount of olefinic hydrocarbonsproduced. With respect to fluidized catalytic cracking operations, spacevelocity is generally measured in terms of weight hourly spaceivelocity, which is defined as the weight of charge per hour per weightof catalyst within the reaction zone. Based upon the raw oil charge, aweight hourly space velocity greater than about 10.0 is preferred,having an upper limit of about 25.0. Where the conversion of thesaturate cracking zone feed is relatively low, a portion of the effluentmaterial may be recycled to effect further conversion to propylene andbutylenes.

The saturate cracking zone requires a catalyst specifically tailored forthe production of the unsaturated light hydrocarbons and crackedgasoline. The catalyst can be selected from a number of known materialsincluding amorphous silica-alumina and zeolitic aluminosilicates, bothof which may contain various catalytic components.

Cracking catalysts suitable for use in the saturate cracking zoneinclude silica-alumina, silica-magnesia, silicazirconia and variouscrystalline aluminosilicates which are characterized as having highcracking activities. The preferred crystalline aluminosilicate crackingcatalyst can be used in admixture with the less active amorphous type,or can be present in substantially pure form. The crystallinealuminosilicate may be naturally-occuring or syntheticallyprepared, andincludes faujasite, mordenite, type A or type U molecular sieves, etc.Whether the catalyst com prises a crystalline aluminosilicate, or anamorphous material, selected metals may be combined therewith by way ofion-exchange or impregnation. Such combined metals include the rareearth metals and alkaline metals, alkaline-earth metals, Group VIIImetals, Group V-B metals, etc. Suitable schemes for effecting thecracking of the saturated liquid stream from the catalytic reformingreaction zone are illustrated in U.S. Pats. 3,161,583 (Cl. 208-164) and3,206,393 (Cl. 208164), although specifically directed toward heaviercharge stocks.

Although producing larger quantities of lighter hydrocarbons, saturatecracking may be effected thermally. Thermal cracking conditions includepressures ranging from about atmospheric to about 500 p.s.i.g. andtemperatures of from about 900 F. to about 1500 F. Preferably, thepresent combination process utilizes catalytic cracking for theconversion of the saturated liquid stream.

ALKYLATION ZONE Although the entire normally gaseous portion of theeffluent from the saturate cracking zone may be introduced into thealkylation zone, the preferred scheme involves separating theseunsaturated gases to provide a butylene concentrate and to recover apropylene concentrate. The latter stream affords a variety of usesincluding the production of C -alkylate or isopropyl alcohol, ashereinbefore set forth, condensation to form isopropylbenzene,polymerization to form tetramer, etc. Preferably, the propyleneconcentrate is utilized as the feed to a hydrolysis unit for theproduction of high octane isopropyl alcohol for the introduction intothe unleaded gasoline pool, or in combination with the butyleneconcentrate for alkylate gasoline production. Similarly, although thesaturated gaseous phases from both the hydrocracking and catalyticreforming reaction zones may be introduced directly into the alkylationzone, in admixture with the unsaturated gaseous phase, a preferredscheme involves separating the saturated gaseous phases to produce abutane concentrate and to recover propane. Probably the most economicuse of the latter is as a blending component for liquefied petroleum gas(LPG). In another preferred embodiment, the butane concentrate, providedby the saturated gas separation, is further separated into a normalbutane concentrate and an isobutane concentrate, the latter beingadmixed with the butylenes and propylene for conversion in thealkylation reaction zone. The normal butane concentrate is subjected toisomerization in order to produce additional quantities of isobutane.

Separation of the unsaturated gaseous phase and the saturated gaseousphases may be accomplished in any suitable manner known to the art.However, it is preferred that the saturated gaseous phase be separatedin a zone distinct from the zone utilized to separate the unsaturatedgaseous phases. Techniques suitable for effecting the separation of thegaseous phases utilize distillation, adsorption, stripping, cryogenicseparations, diffusion, etc.

The alkylation reaction zone may be any acidic catalyst reaction systemsuch as a hydrogen fluoride-catalyzed system, or one which utilizes aboron halide in a fixed-bed reaction system. Hydrogen fluoridealkylation is particularly preferred, and may be conducted substantiallyas set forth in U.S. Pat. No. 3,249,650 (Cl. 260-68348). Briefly, thealkylation reaction when conducted in the presence of hydrogen fluoridecatalyst, is such that the catalyst to hydrocarbon wolume ratio withinthe alkylation reaction zone is from about 0.5 to about 2.5. Ordinarily,anhydrous hydrogen fluoride will be charged to the alkylation system asfresh catalyst; however, it is possible to utilize hydrogen fluoridecontaining as much as 10.0% water or more. Excessive dilution with wateris generally to be avoided since it tends to reduce the alkylatin'gactivity of the catalyst and further introduces corrosion problems. Inorde r to reduce the tendency of the olefinic portion of the chargestock to undergo polymerization prior to alkylation, the molarproportion of isoparaffins to olefinic hydrocarbons in an alkylationreactor is desirably maintained at a value greater than 1.0, andpreferably from about 3.0 to about 15.0. Alkylation reaction conditions,as catalyzed by hydrogen fluoride, include a temperature of from 0 toabout 200 iF., and preferably from about 30 F. to about F. The pressuremaintained within the alkylation system is ordinarily at a levelsufficient to maintain the hydrocarbons and catalyst in a substantiallyliquid phase; that is, from about atmospheric to about 40 atmospheres.The contact time within the alkylation reaction zone is convenientlyexpressed in terms of spacetime, being defined as the volume of catalystwithin the reactor contact zone divided by the volume rate per minute ofhydrocarbon reactants charged to the zone.

Usually the space-time will be less than 30 minutes and preferably lessthan about 15 minutes.

When the alkylation reaction of the combination process of the presentinvention is effected as a fixed-bed unit utilizing a boron halidecatalyst, such as boron trifluoride, the operating conditions will besubstantially those as set forth in U.S. Pat. No. 3,200,165 (Cl.260-671), notwithstanding that this patent teaches analkylation-transalkylation process for the production of ethylbenzene.The amount of boron trifluoride is relatively small, and generally notmore than about 1.0 gram of boron trifluoride per gram mole of theolefinic hydrocarbon. The boron trifluoride alkylation reaction zone isof the conventional fixed-bed type, and contains a borontrifiuoride-modified inorganic oxide selected from diverse inorganicoxides including alumina, silica, boria, oxides of phosphorous, titania,zirconia, zinc oxide, mixtures of two or more, etc. The operatingconditions may be varied over a relatively wide range, the preciseselection generally being dependent upon the character of both theolefinic hydrocarbon as well as the paraffinic hydrocarbon. In anyevent, the alkylation reaction may be effected at a temperature of aboutto about 250 C., and under a pressure of from about 15 to about 200atmospheres, or more. The pressure is usually selected to maintain thereaction mixture substantially in the liquid phase at the selectedoperating temperature. The liquid hourly space velocity through thealkylation reaction zone may likewise be varied over a relatively widerange of from about 0.1 to about 20.0, or more. In addition to theforegoing, suitable alkylation reaction schemes are indicated in U.S.Pats. 2,832,812 (Cl. 260483.42) and 2,818,415 (Cl. 260683.4).

ISOMERIZATION ZONE As hereinbefore set forth, a preferred schemeinvolves separating a saturated gaseous phase from both thehydrocracking and catalytic reforming reaction zone to provide a butaneconcentrate. In addition to the butane concentrate, the saturated gasseparation zone permits recovery of a hydrogen-rich stream which may berecycled to the hydrocracking and reforming zones, a methaneethaneconcentrate which is utilizable as fuel gas and a propane concentratefor use as LPG. With respect to the butane concentrate, although it maybe introduced directly into the alkylation reaction zone, a preferredscheme involves further separation to concentrate the isobutane foralkylation purposes, and the normal butanes which serve as the chargestock to the isomerization zone. As indicated in U.S. Pat. No. 2,900,425(Cl. 260-666), the isomerization process is effected in a fixed-bedsystem utilizing a catalytic composite of a refractory inorganic oxidecarrier material, a Group VIII noble metal component and a metal halideof the Friedel-Crafts type. As previously indicated, the refractoryoxide carrier material may be selected from the group of metallic oxidesincluding alumina, silica, titania, zirconia, alumina-boria,silicazirconia, and various naturally-occurring refractory oxides. Ofthese, a synthetically-prepared gamma alumina is preferred. The GroupVIII noble metal is generally present in an amount of about 0.01% toabout 2.0% by weight, and may be one or more metals selected from thegroup of ruthenium, rhodium, osmium, iridium, and particularly platinumor palladium. Suitable metal halides of the Friedel-Crafts type includealuminum chloride, aluminum bromide, ferric chloride, ferric bromide,zinc chloride, beryllium chloride, gallium chloride, titaniumtetrachloride, zirconium chloride, stannic chloride, etc. The quantityof the Friedel-Crafts metal halide will be within the range of about2.0% to about 25.0% by weight.

The isomerization reaction is preferably effected in a hydrogenatmosphere utilizing suificient hydrogen so that the hydrogen tohydrocarbon mole ratio of the reaction zone feed will be Within therange of from about 0.25 to about 10.0. Operating conditions willfurther include temperatures ranging from about 100 C. to about 300 C.,although temperatures within the more limited range of about C. to about275 C. will generally be uti lized. The pressure under which thereaction zone is maintained will range from about 50 to about 1,500p.s.i.g. A fixed-bed type process is preferred, with the butane andhydrogen feed passing through the catalyst in downward flow. Thereaction products are separated from the hydrogen, which is recycled.and subjected to fractionation and separation to produce the desiredreaction product. Recovered starting material is also recycled so thatthe overall process yield is high. Liquid hourly space velocities willbe maintained within the range of about 0.25 to about 10.0, andpreferably within the range of about 0.5 to about 5.0. Another suitableisomerization process, for the production of isobutane, is found in U.S.Pat. No. 2,924,628 (Cl. 260-666).

AROMATIC EXTRACTION ZONE As hereinbefore set forth, the catalyticreforming zone is maintained under relatively low severity operatingconditions in order to produce a product efiluent rich in aromatichydrocarbons and normally liquid saturates including paraflins andcyclo-paraffins. The saturated material is converted into a crackedgasoline in the saturate cracking zone. Although the entire liquidportion of the reforming zone efiluent, including the aromatics, may beintroduced into the saturate cracking zone, a preferred techniqueinvolves separating the reformed product effiuent to recover thearomatics contained therein. Although any separation scheme may beutilized, a greater degree of efficiency is achieved through the use ofa solvent extraction zone. Solvent extraction to produce an aromaticconcentrate and a paraffinic rafiinate is a well known technique whichis thoroughly described in the literature. Suitable techniques involvethe operations illustrated in U.S. Pats. Nos. 2,730,558 (Cl. 260-674)and 3,361,664 (Cl. 208313).

The solvent extraction process utilizes a solvent having a greaterselectivity and solvency for the aromatic components than for theparaffinic components contained in the reformed product effluent.Selective solvents may be selected from a Wide variety of normallyliquid organic compounds of generally polar character; that is, compounds containing a polar radical. The selective solvent is one whichboils at a temperature above the boiling point of the hydrocarbonmixture at the ambient extraction pressure. Illustrative specificorganic compounds, useful as selective solvents in extraction processesfor the recovery of aromatic hydrocarbons include the alcohols, such asmethanol, ethanol and higher homologous monohydric alcohols; theglycols, such as ethylene glycol, propylene glycol, butylene glycol,tetra-ethylene glycol, glycerol, etc; the glycol ethers, such asdi-ethylene glycol, di-propylene glycol, di-methyl ether of ethyleneglycol, tri-ethylene glycol, tri-propylene, etc.; other organic solventswell known in the art for extraction of hydrocarbon components frommixtures thereof with other hydrocarbons may be employed. A particularlypreferred class of solvents are those characterized as thesulfolanetype. Thus, as indicated in U.S. Pat. No. 3,470,087 (Cl.208-321), the preferred solvent is one having the S-membered ring, oneatom of which is sulfur, the other four being carbon and having twooxygen atoms bonded to the sulfur atom. In addition to sulfolane, thepreferred class includes the sulfolenes such as 2-sulfolene and3-sulfolene.

The aromatic selectivity of the preferred solvents can be furtherenhanced by the addition of water. This increases the selectivity of thesolvent phase for aromatic hydrocarbons over non-aromatic hydrocarbonswithout reducing substantially the solubility of the solvent phase foraromatic hydrocarbons. The solvent composition contains from about 0.5%to about 20.0% by weight of water, and preferably from about 2.0% toabout 15.0%, depending on the particular solvent and the processconditions under which the extraction, extractive distillation andsolvent recovery zones are operated.

In general solvent extraction is conducted at elevated temperatures andpressures selected to maintain the charge stock and solvent in theliquid phase. Suitable temperatures are within the range of from 80 F.to about 400 F. and preferably from about 150 F. to about 300 F.Operating pressures include super-atmospheric pressures up to about 400p.s.i.g., and preferably from about 50 p.s.i.g. to about 150 p.s.i.g.

Typical extractive distillation zone pressures are from atmosphericpressure up to about 100 p.s.i.g., although the pressure at the top ofthe distillation zone will generally be maintained in the range of about1 p.s.i.g. to about 20 p.s.i.g. The reboiler temperature is dependentupon the composition of the feed stock and the selected solvent,although temperatures of from about 275 F. to 360 F. appear to yieldsatisfactory results. The solvent recovery system is operated at lowpressures and sufiiciently high temperatures to drive the aromatichydrocarbons overhead, thus producing a lean solvent bottoms stream.Preferably, the top of the solvent recovery zone is maintained atpressures of from about 100 to about 400 millimeters of mercuryabsolute. These low pressures must be used since the reboilertemperature should be maintained below about 370 F. to avoid thermaldecomposition of the organic solvent.

DESCRIPTION OF DRAWING A preferred embodiment of the present inventionis illustrated in the accompanying drawing. The illustration ispresented by Way of a block-type flow diagram, in which each blockrepresents one particular step or stage of the process. Miscellaneousappurtenances, not believed necessary for a clear understanding of thepresent combination process, have been eliminated from the drawing. Theuse of details such as pumps, compressors, instrumentation and controls,heat-recovery circuits, miscellaneous valving, start-up lines andsimilar hardware, etc., is Well within the purview of one skilled in theart. Similarly, with respect to the flow of materials throughout thesystem, only those major streams required to illustrate theinterconnection and interaction of the various zones are presented.Thus, various recycle lines, vent gas streams, etc., have beeneliminated also.

The drawing will be described in conjunction with a commercially-scaledunit designed to process 100,000 bbl./day of a full boiling range crudeoil having a gravity of about 39.4 API and containing about 0.3% byweight of sulfurous compounds, calculated as elemental sulfur. From anexternal source, normal butane, in an amount of 5,000 bbl./day, andisobutane in an amount of about 2,000 bbl./ day is provided. Notillustrated in the drawing is the initial separation of the crude oil byway of distillation techniques and vacuum separation which results in5,400 bbl./day of an asphaltic residuum, 5.372 bbl./day of a low sulfur,No. 6 bunker fuel, 2,772 bbl./day of a No. 2 fuel oil and 3,100 bbl./dayof a range oil. Of the remaining 83,356 bbL/day, 53,119 constitutes thecharge stock to the hydrocracking zone of the present process, and30,237 bbL/day constitutes the crude oil distillation overhead. Thisoverhead stream contains those hydrocarbons boiling below a temperatureof about 400 F., and is combined with the 7,000 bbl./day of externalnormal butane and isobutane.

The overhead stream from the crude distillation unit is initiallyintroduced into a debutanizer from which a C -minus stream, containing aminor quantity of pentanes, is removed in an amount of about 10 ,606bbL/day. A bottoms stream, principally C -400 F. hydrocarbons. is passedinto a splitter in order to recover 21,993 bbl./day of a heptane-400 F.naphtha fraction to serve as a portion of the charge to the catalyticreforming zone. The C /C concentrate, in an amount of 4,638 bbl./ day,forms part of the feed to the saturate cracking zone. Since the intendedcharge to the reforming reaction zone contains detrimental quantities ofsulfurous and nitrogenous compounds, it is first introduced into anaphtha hydrorefining reaction zone wherein the contaminating influencesare converted into hydrogen sulfide, ammonia and hydrocarbons. Theintended object of this commercially-scaled unit is to produce anunleaded gasoline pool having an octane rating (Research Method) of atleast 95.0.

Referring now to the accompanying drawing, the 53,119 bbl./day of 400'F.-plus material from the crude distillation column is introduced intohydrocracking zone 2 by way of line 1. In the present illustration,hydrocracking zone 2 is a two-stage system functioning in series flow,in the first stage of which sulfurous and nitrogenous compounds areconverted into hydrogen sulfide and hydrocarbons, and some conversioninto lower-boiling hydrocarbons is effected. The second stage involvesadditional hydrocracking into gasoline boiling range materials. Thecatalytic composite disposed in the first stage of hydrocracking zone 2is a composite of 1.8% by weight of nickel and 16.0% by weight ofmolybdenum combined with an amorphous alumina-silica carrier materialcomprising 37.0% by weight of silica. The charge is admixed with about10,000 scf./bbl. of hydrogen, and passes into the reaction zone at atemperature of about 650 F., a pressure of about 2,100 p.s.i.g. and aliquid hourly space velocity of about 1.63. An increasing temperaturegradient of F. is experienced, and the product efiiuent is withdrawn ata temperature of about 750 F. Following its use as a heat-exchangemedium, to reduce the temperature to about 700 F., the product effluent,without intermediate separation, passes into the second stage whichcontains a catalytic composite of 5.0% by weight of nickel combined witha crystalline aluminosilicate carrier material, about 93.0% by weight ofwhich is zeolitic and constitutes faujasite. The operating pressure isabout 2,050 p.s.i.g., the fresh feed liquid hourly space velocity isabout 0.90 and about 4,000 scf./bbl. (on fresh feed) of hydrogen isutilized as a quench stream to maintain the increasing temperaturegradient at a level of about 20 F.

The effluent is separated, following use as a heatexchange medium andfurther cooling to a temperature of about 100 F., in a high-pressure(2,000 p.s.i.g.) separator to provide a hydrogen-rich gaseous phase anda principally liquid phase. The gaseous phase may be subjected to one ormore treatments to remove hydrogen sulfide and light, saturatedhydrocarbons to increase the hydrogen concentration prior to recyclethereof to the first stage of the system. The principally liquid phase1s separated into various component streams, and hydrocarbonaceousmaterial boiling above 400 F. is recycled to combine with the firststage effluent, thereby providing a combined liquid feed ratio to thesecond stage of about 1.60. Overall hydrogen consumption is 3.48% byweight, or 2,117 scf./bbl., based upon fresh feed to the first stage.Th1; firlial product distribution and yield is presented in TABLE IHydrocracking Yields and Distribution Weight percent Volume percentComponent Ammonia Hydrogen sulfide Methane Pentanelhexanm Heptane-400 Fmixture with the hydrorefined straight-run naphtha stream from the crudedistillation column in line 4; the total liquid naphtha feed toreforming zone 5 is 61,106 bbl./ day, of which 41,007 bb1./day is thehydrocracked naphtha from line 3. The pentane/hexane hydrocracked lightnaphtha concentrate, in an amount of 16,052 bbl./ day may be included aspart of the liquid feed to reforming zone 5. However, since it has aresearch octane rating (RON) of about 84.7, a preferred technique, notillustrated in the accompanying drawing, recovers this stream as aseparate product from hydrocracking zone 2, and incorporates the sameinto unleaded gasoline pool 13.

Catalytic reforming zone 5, in the present illustration, is designed toproduce maximum quantities of a normally liquid reformate product havingan RON of about 92.0. This product stream is intended to includepentanes and higher-boiling hydrocarbons in the gasoline boiling range.The reforming reaction zone 5 contains a catalytic composite of alumina,0.375% by weight of platinum, about 0.25% by weight of germanium andabout 0.85% chloride, calculated as the elements. The operatingconditions imposed upon the catalytic reforming zone include a pressureof 250 p.s.i.g., a liquid hourly space velocity of about 2.3, an averagecatalyst bed temperature of about 940 F. to about 970 F. and a hydrogento hydrocarbon mole ratio of about 5.0: 1.0.

The reformed product efiiuent is separated within reforming zone 5 intoa hydrogen-rich recycle stream, a butane-minus gaseous phase in line 6and a pentanecontaining, normally liquid stream in line 9. Catalyticreforming is a hydrogen-producing process, and excess h'ydrogen overthat required to satisfy the mole ratio of 5.0: 1.0 may be employed tosupplant that hydrogen consumed within hydrocracking reaction zone 2.The saturated gaseous phase in line 6 is admixed with the saturatedgaseous phase in line 7, the mixture continuing through line 6 intosaturated gas separation zone 8. The 10,606 bbl./day of C -minusconcentrate, including the 7,000 bbl./day of outside butanes, is alsointroduced into saturated gas separation zone 8. The pentane-plusportion of the reformed product effluent preferably passes through line9, in an amount of about 50,494 bbl./day, into extraction zone 10.

As hereinbefore set forth, extraction zone 10 is not consideredessential to the present combinative process, but constitutes onepreferred embodiment. The normally li'quid, catalytically reformedproduct effiuent is introduced into the lower portion of an extractioncolumn countercurrently to a lean solvent stream introduced into anupper portion of said column, the mole ratio of solvent to hydrocarbonbeing about 3.2:1.0. The selected solvent is sulfolane, and the columnconditions are a top pressure of about p.s.i.g. and a reboilertemperature of about 320 F. A saturate-rich raflinate stream iswithdrawn as an overhead product while the rich solvent bottoms streamis introduced into an extractive distillation zone. Additional rafiinateis withdrawn as an overhead stream, com bined with the saturate-richrafiinate from the extraction column, and passed from extraction zone10, via line 11 into saturate cracking zone 14; the saturated raffinateis in an amount of about 23,484 bbL/day. \Rich solvent is introducedinto a solvent recovery zone functioning at sufiiciently low pressuresand high temperatures to drive aromatic h'ydrocarbons overhead whileproducing a lean solvent bottoms stream for recycle to the extractioncolumn. The aromatic concentrate, in'an amount of about 27,010 bbL/day,is withdrawn from extraction zone 10 through line 12 and introducedthereby into unleaded gasoline pool 13. With respect to the rafiinatestream in line 11, beneficial results are obtained, with respect tosubsequent saturate cracking, when the stream is substantiallysulfolane-free. One possible technique for removing solvent from thesaturate-rich raffinate is that disclosed in US. Pat. No. 3,470,087 (Cl.208-321).

Including the 4,638 bbl./day of C /C light naphtha from the crudedistillation column, the total feed to saturate cracking zone 14 is28,122 bbL/day. The cracking of the saturated rafiinate is effected in afluid catalytic cracking system utilizing a cracking catalyst containing10.0% by weight of faujasite in a silica matrix and about 2.7% by weightof alumina. The system functions much the same as the well known fluidcatalytic cracking process which is exemplified in US. Pats. 3,161,583(Cl. 208- l 64) and 3,206,393 (Cl. 208-164). The saturate cracking iseffected at a reactor temperature of about 1050 F., the regeneratortemperature being about 1250 F., and with a catalyst to charge stockweight ratio of 15.6:1.0. Since there is no liquid recycle to thereactor, the combined liquid feed ratio is 1.0. The results of thesaturate cracking operation are presented in the following Table II:

TABLE II saturate cracking yields and distribution Weight percent VolumeComponent percent Ethylene- Propane Propylene- Butylenes Pentane-400 Thepentane-400 F. gasoline fraction, in an amount of 9,247 bbl./day, isremoved from saturate cracking zone 14 through line 16, and, via line17, is sent to the unleaded gasoline pool 13. An unsaturated gaseousphase, comprising butylenes and lighter components, being about 60.0% byweight of the feed to saturate cracking zone 14, passes by way of line15 into unsaturated gas separation zone 18. With respect to this gaseousphase, the hydrogen, methane, ethane and ethylene may be first separatedin the saturate cracking zone so that the feed to gas separation zone 18is primarily a C /C -concentrate rich in propylene and butylenes. Theparticular means selected for separating the gaseous phase into thedesired component streams is not pertinent to my invention, and anytechniques described in the art, and easily recognized by those havingexpertise therein, may be utilized.

In the commercial design being used to illustrate this embodiment, thequantity of propylene is such that about 500 bbl./ day are withdrawnfrom the process through line 26, for use, in this instance, as a rawmaterial for the production of isopropyl benzene. The remainder of thepropylene is withdrawn through line 19 in admixture with the butanes andbutylenes, and introduced via line 20 into al-kylation zone 24. In thisfashion, both C -alkylate and C -alklylate gasolines are produced.

Saturated gas separation zone 8 serves to concentrate theC.,-hydrocarbons contained in (1) the gaseous phase from the crudedistillation column, (2) the saturated vaporous phase in line 7 fromhydrocracking reaction zone 2 and (3) the saturated vaporous phase inline 6 from catalytic reforming zone 5. A propane concentrate is removedfrom the process system by way of line 25, in an amount of about 57,234lbs./hr., or about 7,725 bb./day. Propane may be dehydrogenated topropylene to produce additional C -alkylate gasoline, tetramer orisopropyl benzene, or utilized as a component of LPG.

Although the total C -hydrocarbon stream may be sent directly toalkylation reaction zone 24, with subsequent recycle of unreactedbutanes to isomerization zone 23, a preferred technique involves furtherseparation into an isobutane concentrate and a normal butane stream. Theformer is removed from separation zone 8 through line 22, and is passedvia lines 19 and 20 into alkylation reaction zone 24; the latter, inline 21, forms part of the feed to isomerization reaction zone 23, theremainder of the feed being recycled unreacted butanes from thealkylation reaction zone 24 (the recycle line is not illustrated in thedrawing).

Isomerization reaction zone 23 utilizes a fixed-bed reaction zonecontaining a catalytic composite of alumina, about 19.0% by weight ofaluminum chloride, and 0.375% by weight of platinum, calculated as theelemental metal. The reaction zone is maintained under a pressure ofabout 300 p.s.i.g., a temperature of about 300 F. and a hydrogen tohydrocarbon molal ratio of about 1.0: 1.0. The reactants traverse thecatalyst bed at a liquid hourly space velocity of about 1.0. Followingseparation of gaseous material, the butane portion of the productefliuent is passed into alkylation reaction zone 24 by way of line 20,in admixture with the butane/butylene concentrate in line 19.

Alkylation reaction zone 24 is a hydrofluoric acid system which producesabout 23,341 bbl./day of alkylate gasoline, of which about 9,381 bbL/dayis C -alkylate. The reaction time, utilizing a pumped acid settlerreactor, is about nine minutes and the acid/hydrocarbon volume ratio isabout 1.5 1.0. The alkylation reactions are effected at a temperature ofabout 100 F. Following separation of unreacted butanes, which arerecycled to isomerization reaction zone 23, and a minor quantity ofcracked products, the alkylate gasoline passes through line 17 intounleaded gasoline pool 13.

A summary of the results of the foregoing combinative process ispresented in the following Table III;

TABLE TIL-UNLEADED GASOLINE POOL 1 The vapor pressure specification ofthe entire gasoline pool requires addition of butanes to reach a levelof 10.0.

From Table III, it is noted that 86,028 bbL/day of a 98.6 RON (researchoctane rating) unleaded gasoline pool is produced by the presentcombination process. Based upon 100,000 bbl./day of total crude oilcharge and 7,000 bbl./day of outside butanes (107,000 bbL/day total),the volumetric yield of gasoline is 80.4%. Economic studies indicatethat the incremental cost of producing this unleaded gasoline pool, withrespect to the commercial unit being considered, is only $0.268/bbl., or0.638 cents/gallon.

The foregoing demonstrates the method by which the present invention iseffected and the benefits afforded through the utilization thereof.

I claim as my invention:

1. A process for producing a high octane, unleaded gasoline pool whichcomprises the steps of:

(a) reacting a heavier-than-gasoline charge stock with hydrogen in acatalytic hydrocracking reaction zone, at hydrocracking conditionsselected to produce gasoline boiling range hydrocarbon products;

(b) separating the resulting hydrocracked product effluent to provide afirst substantially saturated vaporous phase and a gasoline boilingrange, normally liquid stream;

(c) reacting at least a portion of said liquid stream and hydrogen in alow-severity catalytic reforming reaction zone, at reforming conditionsselected to convert naphthenic hydrocarbons into aromatic hydrocarbons;

(d) separating the resulting reformed product effluent to provide anaromatic concentrate, a saturated normally liquid stream and a secondsubstantially saturated vaporous phase;

(e) reacting at least a portion of said saturated normally liquid streamin a saturate cracking reaction zone, at cracking conditions selected toproduce a cracked gasoline boiling range liquid stream and asubstantially unsaturated vaporous phase;

(f) reacting at least a portion of said unsaturated vaporous phase Withat least a portion of said first and second saturated vaporous phases inan alkylation reaction zone, at alkylating conditions selected toproduce an alkylate gasoline boiling range, normally liquid stream; and,

(g) recovering said aromatic concentrate, said cracked gasoline streamand said alkylate gasoline stream as said high octane, unleaded gasolinepool.

2. The process of claim 1 further characterized in that said first andsecond saturated vaporous phases are separated to provide a butaneconcentrate, at least a portion of which is reacted with at least aportion of said unsaturated vaporous phase in said alkylation reactionzone.

3. The process of claim 2 further characterized in that said butaneconcentrate is separated to provide a normal butane concentrate and anisobutane concentrate, and said butane concentrate is reacted withhydrogen in a hydroisomerization reaction zone, at isomerizingconditions selected to product isobutane isomers.

4. The process of claim 3 further characterized in that said isobutaneisomers and said isobutane concentrate are reacted in said alkylationzone with at least a portion said unsaturated vaporous phase.

5. The process of claim 1 further characterized in that said unsaturatedvaporous phase is separated to provide a propylene concentrate and abutylene concentrate, the latter being reacted, at least in part, withat least a portion of said first and second vaporous phases in saidalkylation reaction zone.

6. The process of claim 4 further characterized in that the alkylationreaction zone efiiuent is separated to recover unreacted butanes, saidbutanes being recycled to said hydroisomerization reaction zone.

7. The process of claim 1 further characterized in that said saturatecracking reaction zone is a catalytic cracking zone.

8. The process of claim 1 further characterized in that said saturatecracking reaction zone is a thermal cracking reaction zone.

9. The process of claim 5 further characterized in that said propyleneconcentrate is converted to an alcohol in a hydrolysis reaction zone.

10. The process of claim 1 further characterized in that said reformedproduct efiiuent is separated in a solvent extraction zone.

References Cited UNITED STATES PATENTS 2,818,459 12/1957 Gautt 260-683482,924,628 2/1960 Donaldson 260-666 R 3,132,087 5/1964 Kelley et al.208-60 3,249,650 5/1966 Fenske 260-68348 3,361,664 1/1968 Broughton etal 208-313 3,513,085 5/1970 Leas 208-60 DELBERT E. GANTZ, PrimaryExaminer G. E. SCHMITKONS, Assistant Examiner US. or. X.R.

